Desulfurizing and reforming naphthas



Nov. 30, 1954 H. G. MCGRATH 2,695,866

DESULFURIZING AND REFORMING NAPHTHAS Filed Dec. 27, 1951 5 Sheets-Sheets FIG.4

WT. l SULFUR IN C -FREE GASOLINE I 2 3 4 5 6 7 8 9 l ll l2 VOLUME NAPHTHA REFORMED PER VOLUME OF CATALYST 95 FIG. z

9 90 h m 2 a E D B5 o o U O 80 I Z 3 4 5 8 7 B 9 l0 II I? w VOLUME NAPHTA REFORMED PER VOLUME OF CATALYST z 90 I O m 3 ou 55 o o I o T 4 FIG. 6 Q U 80 l 0 o 75 :1 l NVENTOR g I HENRY G. Mc GRATH 9' 7o M 4O 5O 60 I00 I20 I40 I60 I 200 220 HYDROGEN PARTIAL PRESSURE,PSIA. ATTORNEYS United States Patent DESULFURIZING AND REFORMING NAPHTHAS Henry G. McGrath, Union, N. 1., assignor to The M. W. Kellogg Company, Jersey City, N. 1., a corporation of Delaware Application December 27, 1951", Serial No. 263,661

6 Claims. (Cl. 196-28) This invention relates to an improved process of reforming and desulfurizing hydrocarbon oils, and more particularly pertains to a-novelproc'ess' for reforming and desulfurizing petroleum naphthas of low octane quality and high sulfur content.

It is an object of this invention to provide an improved process for reforming and desulfurizing highsulfurcontaining hydrocarbon oils.

Another object of thisinvention is to provide an improved process for reforming and desulfurizing high sulfur-containing naphtha stocks.

Still another object of this invention is to provide an improved process forreforming and desulfurizing a high sulfur-containing cracked petroleum naphtha fraction.

Other objects and advantages of this invention will become apparent from the following description and explanation thereof.

In accordance with the present invention an improved process is provided which comprises reformingand desulfurizing a sulfur-containing hydrocarbon oil by contacting same with a catalyst selected from the group consisting of the oxides of chromium and molybdenum, the relative proportions-being not greater than about 7 volumes of naphtha per volume of catalyst, in the presence of hydrogen under suitable conditions of temperature and pressure.

Another aspect of the present invention is concerned with the process of desulfurizing and reforming a sulfur containing hydrocarbon oil by maintaining a hydrogen partial pressure of at least about 65 p. s. i. a.

The hydrocarbon oils processed in accordance with this invention include, generally, naphthas which have an initial boiling point in the range of about 90 to about 300 F., and an endpoint in the range of about 350 to about 500 F. The naphthas are of low octane quality, and generally, they contain a high sulfur content in the order of about 0.3 to about 2.5% byweight. Such stocks are derived from West Texas oil fields and those situated along the Pacific Coast, which are known to be of high sulfur content. Naphtha can be a straight run or'virgin stock, one which is cracked either thermally or catalytically, or a blend of straight run or virgin and cracked stocks.

Naphthas containing a large amount of sulfur and which are of low octane quality present a difficult problem of reforming in view'that it is desired to effectan octane improvement as well as remove substantially all of the sulfur which is present in the material. According to present day technology, effective methods are available for processing such types of stocks for either desulfurization or reforming, but as yet thereis lacking a highly efficient method for both reforming and desulfurizing suchv types of naphthas. In the present instance, it was found that a naphtha containing an unusually high amount of sulfur can be processed satisfactorily by employing conditions which will result in a net consumption of hydrogen. Furthermore, as will be shown later, under conditions of net consumption of hydrogen, it was quite unexpectedly, discoveredthat unusual advantages are obtained whenprocessing not more than about 7 volumes of naphthaper volume of catalyst.

The catalytic material which is employed for this invention includes the oxides of chromium and/ or molybdenum. These catalyticagents can be used alone or they can be supported on a carrier material, suchas for example,. alumina, activated carbon, silica, magnesia, silicaalumina, silica-magnesia, fullers earth, kieselguhr,

"ice

pumice, etc. The catalytic agents, i. e., the oxides of chromium and/or molybdenum constitute about 0.1 to about 25% by weight'of the total catalyst. The oxide of chromium is preferably employed in the amount of about 5 to about 25% by weight, based on the total catalyst; whereas the oxide of molybdenum is preferably used in the amount of about 5 to about 1.5% by weight, based on the total catalyst. According to some methods of catalyst preparation, molybdenum trioxide is formed. It is preferred to employ the lower form of oxide of molybdenum as the catalytic agent, because substantially improved results with respect to reforming and desulfurizing are obtained. Hence, it is desirable to subject molybdenum trioxide to a hydrogen treatment prior to use under reforming conditions. In this manner, the trioxide is reduced to the lower form of oxide, e. g., the dioxide, and it is believed by some workers that molybdenum is also produced by this treatment. Generally, the reduction treatment involves contacting the catalyst with about 0.3 to about 10 cubic feet of hydrogen (at 60 F. and 760 mm.) per pound of catalyst at a temperature of about 800 to about 1200 F., and a pressure of about 50 to about 500 p; s. i. g. In a fixed bed operation, the reduction treatment is effected by passing a hydrogen-containmg gas over the catalytic material after it has been regenerated by means of an oxygen-containing gas. In a moving bed system, the reduction is accomplished by passing the regenerated catalyst from the regeneration zone to a separate reduction zone prior to being returned to the r'eactionzone.

Other catalysts which can be used for reforming-desulfurization of hydrocarbon oils are, for example, tungsten sulfide, mixed nickeland tungsten sulfides, cobalt-molybdena on alumina, etc.

The reforming and desulfurization procedure is effected as either a fixed bed or moving bed operation in which a fluid or non-fluid technique is employed. For any type of operation, the catalyst can be employed as either a lump, granular, pelleted or finely divided material. In the granular state, the particles can have a size in the range of about 0.1 to about 25 mm.; whereas the finely divided material can have a particle size in the range of about 5 to 250 microns, or more usually, about 10 to about microns. As a non-fluid or static bed, the catalyst is supported within the reaction zone and the reactant materials are passed either upwardly or downwardly between the particles and through the pores in the particles. By passing the reactant materials ina downward direction through the catalystmass, there is a tendency to avoid any lifting of catalyst mass which causes an undue vibration of apparatus or equipment parts.

As a result of contacting the catalytic material with the reactants, there is a deposition of carbonaceous material on the catalyst which causes a temporary deactivation. In order to regenerate-the catalyst, the passage of the reactants may be discontinued and the catalytic material, with or without suitable purging by means of steam, etc., is then contacted with a regeneration gas or an oxygencontaining gas, e. g., air or oxygen. This type of an operation is cyclic, in that, the process phase is conducted for a definite period, followed by a regeneration phase for a given period of time. Ordinarily, for a fixed bed operation, the reaction phase or period is conducted for about 0.25 to about 24 hours, preferably for about 0.5 to about 15 hours. The regeneration follows in alternate cycles and it may take about 0.5 to about 20 hours, and preferablyabout 1 to about 10 hours. In order to maintain a continuous flow of processing materials, there should be at least two processing vessels, so that while the catalyst in one vessel is undergoing regeneration, the other vessel is in the reaction cycle. Therefore, a fixed bed operation may be continuous in the sense that processing materials are notinterrupted. However, it requires the use of at least two vessels for such a purpose.

The present invention can be practiced as a fluid system. In this type of an operation, the reactant materials are passed upwardly through a mass of catalyst at a velocity sufiicient to cause-fiuidization. Generally, the upfiowing fluids shouldcontain a superficial linear gas velocity in the order of about 0.1 to about 50 feet per second, or more usually, about 0.1 to about 6 feet per second. In this range of gas velocities, fluidization occurs to produce either a lean or a dense phase of catalyst. The gas velocity employed will depend upon the density and size of the catalyst particles. Dense and large particles require higher gas velocities than the smaller and less dense particles. For most purposes, it is preferred to employ a dense phase, because a more intimate contact between gases or vapors and catalyst particles is effected thereby. In the moving bed-fluid system, the reaction is conducted in a separate reaction zone. There is a continuous withdrawal of catalyst from the reaction zone, and this material is then passed to a regeneration zone, wherein it is contacted with an oxygen-containing gas. Hence, there is a continuous flow of processing materials, i. e., reactants and regeneration gas, as well as a continuous flow of catalyst from the reactor to the regenerator and vice versa. As a result, there is a tendency for uniform mixing of catalyst and catalyst activity to exist in each zone, thereby tending to avoid temperature gradients, uneven distribution of regenerated catalyst, poor product distribution, etc.

In the practice of this invention, it is important that not more than 7 volumes of naphtha per volume of catalyst be processed. For a fixed bed operation, this condition can be readily measured as the volume of naphtha which is charged for a given reaction cycle per volume of catalyst which is present in the vessel. As for a moving bed operation, whether it be fluid or non-fluid, this condition is measured as the volume of naphtha charged to the reactor, on a given time basis, per volume of regenerated catalyst which is circulated to the reactor on the same time basis. In other words, for a moving bed operation, the condition is measured as the inverse of the catalyst-oil ratio, on a volumetric basis. It will be shown hereinafter that the processing of not more than 7 volumes of naphtha per volume of catalyst results in unusual ad vantages with respect to desulfurization and reforming. In the operation of my process generally, a temperature of about 750 to about 1150 F. is employed, while maintaining the pressure in the range of about 50 to about 300 p. s. i. g. Under the conditions specified above, the volumetric space velocity can be about 0.1 to about 4. The volumetric space velocity is measured as the cubic feet of oil charged per hour per cubic foot of catalyst present in the reactor. In order to effect the consumption of hydrogen, it is necessary to maintain a hydrogen partial pressure of at least about 50 p. s. i. a. This condition can be employed for both an oxide of molybdenum and an oxide of chromium. However, in view that there are differences in the catalytic properties of these materials, the conditions of operationwill vary in order to obtain the best performance for each type of catalyst. With respect to a molybdena catalyst or an oxide of molybdenum, it is preferred to employ a temperature of about 850 to about 1050 F. while maintaining a reaction pressure of about 75 to about'250 p. s. i. g. The preferred hydrogen partial pressure is at least about 65 p. s. i. a. and it can be varied up to about 200 p. s. i. a. or higher. The volumetric space velocity is about 0.2 to 2.0. In the case of an oxide of chromium, it is preferred to employ a temperature of about 900 to about 1025 F. while maintaining a reaction pressure of about 50 to about 200 p. s. i. g. The hydrogen partial pressure is preferably at least about 65 p. s. i. a.; whereas the volumetric space velocity is about 0.1 to about 1.5.

With respect to desulfurization the notable difference between the oxide of chromium and the oxide of moylbdenum is that the chromium oxide tends at all times to remove sulfur as hydrogen sulfide. In the case of the molybdenum oxide, the sulfur is removed by either combining with the catalytic material and then being burned off as a sulfur oxide during the regeneration step, and/ or it is removed as hydrogen sulfide. The preponderance of a particular method of sulfur removal in the case of an oxide of molybdenum depends upon the operating conditions. In this regard, it is noted that at higher hydrogen partial pressures, there is more of'a tendency for the sulfur to be removed as hydrogen sulfide. Thus in a series of tests conducted at hydrogen partial pressures of 150-200 p. s. i. a., the bulk of the sulfur was removed as HzS. However, at a hydrogen partial pressure of less than 70 p. s. i. a., the bulk of the sulfur will normally be removed with the catalytic material and burned off during regeneration. It is noted, however, that chromium oxide does produce higher yields of reformed gasoline and lower carbon yields than molybdenum oxide for a given octane level of product. This represents a distinct advantage in view that lower carbon yields reduce operating costs and a higher reformed gasoline yield represents a lower cost for manufacturing a given volume of desulfurized reformed gasoline.

For both types of catalysts, regeneration is conducted in a similar manner. The catalysts are contacted with an oxygen-containing gas, e. g., air or oxygen, at a temperature of about 650 to about 1250 F. Generally, the catalysts can be permitted to acquire a carbon content of about 1 to about 20% by weight, before being subjected to a regeneration treatment. A carbon content higher than the range specified is not desired, because there is a more than desired tendency for the quality of the product to be adversely affected due to the lowered catalyst activity. When excessive catalyst deposits accumulate, problems also arise with respect to satisfactory regeneration.

In order to more fully describe my invention, reference will be had to the drawings which form a part of this specification.

Figure 1 is a diagrammatical sketch of a fluid system;

Figure 2 is a schematic diagram of a moving bed operatlon;

Figure 3 is a schematic diagram of a fixed bed operation;

Figures 4 and 5 are correlations which indicate the unusual advantages in reforming not more than 7 volumes of naphtha per volume of catalyst; and

Figure 6 is a correlation illustrating the advantage in gnerating at a hydrogen partial pressure of at least about 5 p. s. i. a.

In Figure 1, reactor 10 is a vertical, cylindrical vessel having a conically shaped bottom 11. Oil feed is charged at the rate of about 2000 bbls. (42 gals.=1 bbl.) per stream day into the bottom part of reactor 10 through inlet 12. The feed may be distributed by means of a circular distributor or any other suitable means along the cross-section of the reaction vessel. This distributor is not shown. The reactor 10 contains a straight side length of 69 feet and an internal diameter of about 6 feet, 3 inches. Situated within the reactor just above the conical section 11 is a grid plate 13. Along the inside wall of the reactor above grid plate 13 extends a circular well or regenerated catalyst inlet 15. This well extends above the grid plate to a height of about 6 feet. Snent catalyst is withdrawn from the catalyst through another circular well 16 which extends from the outside of the conical section 11 of the reactor to a height of about 1 foot above the grid plate, inside the reactor. The spent catalyst which flows into the well 16 is conveyed upwardly in a spent catalyst circular riser 17 whose internal diameter is about 2 inches. The catalyst is conveyed upwardly in riser 17 by means of a gas which flows through a vertical hollow conduit 18 which is disposed co-axially with the riser. The conduit 18 may or may not be of a movable type. The rate of solids flowing through the riser may be regulated by adjusting the rate of gas flow through the conduit 18 and/ or by adjusting the clearance between the bottom of riser 17 and the top of the conduit.

In the reactor, chromia (20% C1'203 on A1203), which is employed as catalyst, may have a fluid density of about 15 to about pounds per cubic foot, or, in the present example, a density of about 40 pounds per cubic foot. To attain this density, the total gas throughput entering at the bottom of the reactor is sufficient to have a superficial linear gas velocity of about 0.1 to about 6 feet per second or, in the present example, a superficial linear gas velocity of about 0.75 feet per second. The amount of catalyst which is present in the reactor will provide a bed depth, under the specified conditions enumerated above, of about 50 feet. Under these conditions, the volumetric space velocity used is about 0.50. By virtue of the temperatures at which the various streams enter the bottom of the reactor, the resultant temperature will be about 950 F. As a result of the reforming and desulfurization of the naphtha, carbonaceous material is produced, which contaminates the catalyst and causes the activity thereof to decline. In the present example, the catalyst of low activity is conveyed upwardly through the riser 17 at a fluid density in the order of about 1 to about 20 pounds per cubic foot. To attain such densities, the superficial linear velocity of the lifting gas should be about 5 to about 80 feet per second. Generally, the density of the solids flownilg upwardly in the riser 17 may be either a lean or dense p ase.

The oil feed is introduced through line 12 at a level above grid plate 13 in order to prevent undesirable thermal cracking of the feed on the hot grid plate. Since the reforming and dehydrogenation reactions are endothermic in nature, it is necessary to supply the heat of reaction in order to obtain the desired rate of conversion. In this respect, a large portion of the heat of reaction is supplied by means of a hydrogen-containing gas or a recycle gas which is introduced into the reactor by means of line 20 at a point below grid plate 13. By reason thereof, the temperature of the grid plate may run, ordinarily, higher than the average temperature of the bed in the bottom section of the reactor. Another consideration is that the regenerated catalyst which enters the reactor through Well 15 will be at a higher temperature than the temperature of the incoming oil feed. Further, the activity of the regenerated catalyst is higher than the average activity of the catalyst in the reactor. Therefore, in order to avoid any undesirable effects which might occur from the regenerated catalyst contacting the oil feed directly, the regenerated catalyst should be introduced at a level above the point of entry of the oil feed. In this manner, the regenerated catalyst has an opportunity to mix with the catalyst already present in the reactor prior to contacting the incoming fresh oil feed. Still another consideration is is position of the spent catalyst well 16, through which spent catalyst is withdrawn from the reactor. As would be expected, the fresh oil feed entering the reactor through line 12 will flow substantially upward. Therefore, by having well 16 below the point of entry of the oil feed, the tendency of oil vapors to be entrained with spent catalyst during the reaction operation is reduced.

Spent catalyst is withdrawn from the reactor through well 16 by means of hydrogen-containing gas or recycle gas which is introduced through line 21. This lifting gas has a temperature of about 700 F. which causes cooling of the spent catalyst. In this manner, less cooling is required in the regeneration step which is to be discussed hereinbelow. The spent catalyst which flows upwardly through riser 17 enters a stripper which is a vertical cylindrical vessel superimposed on reactor 10, but offset from the center line thereof. This stripper contains a lower section 22 which is about 11 feet long and has an internal diameter of about 12 inches. This lower section is joined by an upper enlarged section 23 which has a length of about 14 feet and has an internal diameter of 21 inches. The riser 17 extends through the lower section 22 of the stripper in a vertical position and terminates at the upper section 23 on the bottom of stripper section 22. A stripping gas, such as for example, steam, is introduced at the rate of about 900 pounds per hour through line 25. This stripping gas is used for a catalyst-oil ratio of about 1.0. However, the stripping gas can vary from about 100 to 3000 pounds per hour for the catalyst used in this example. The stripped solids leave the stripper through a segmental well 26 which is located in the bottom part of the lower section 22 and then through a spent catalyst transfer line or standpipe 28, the upper end of which is connected to the well. Gases in the stripping zone leave the stripper through an overhead line 29 and are returned to the top of the reactor. In this example, the pressure in the top of the reactor is maintained at 100 p. s. i. g. By means of a control valve 30, which is installed in line 29, it is possible to regulate the pressure in the stripper above the pressure existing in the reactor. The pressure in the stripper can be maintained up to about 5 pounds p. s. i. g. higher than what exists in the reactor.

In the upper part of the reactor the gases break contact with the fluid bed shown as level 32 and thence flow intoa cyclone separator 33 wherein a further separation of gases from solids is effected. The separated gases leavethe top of the reactor through a product line 35 and are passed to a recovery system which is not shown. The separated solids in the cyclone 33 are returned to the reactor through a long drainpipe or dipleg 36. The dipleg is about 40 feet in length and discharges the separated catalyst into the reactor bed at a point about 10 feet above the grid plate. By having the dipleg extend into the reactor bed to this extent, it is possible substantially to overcome any tendency of the catalyst to classify, that is, the small particles will accumulate in the upper part of the reactor bed and the particles of larger size will ac cumulate in the lower part of the reactor bed.

The stripped solids in standpipe 28 are in a fluidized state and have a density of about 15 to about 75 pounds per cubic foot. In this example, the stripped catalyst is withdrawn from the stripper at the rate of 24,000 pounds per hour and will have a density of about 50 pounds per cubic foot in standpipe 28. This standpipe has a 4 inch diameter and the solids flow therein at the rate of about 3 feet per second. Steam is introduced into standpipe 28 through an aeration line 38 for the purpose of maintaining the flowing solids in a fluidized condition. The rate of the aeration steam through line 38 is about 10 pounds per hour and 400 p. s. i. g. steam. The rate of catalyst through standpipe 28 is automatically controlled by a slide valve 39.

The spent catalyst leaves the standpipe 28 and enters a reactor vessel. This reactor is a vertical cylindrical vessel having an effective internal diameter of about 2 feet 3 inches. This diameter takes into account any insulating material which lines the inside wall of the reactor. The inside overall length is 25 feet. Inside the reactor a catalyst bed having a level 40 is present. This catalyst bed can have a density of about 15 to about 100 pounds per cubic foot. In this example, the density is about 40 pounds per cubic foot. The catalyst is regenerated by burning any contaminating deposits with air which is introduced through a line 41 which is connected to the bottom of the regenerator. The air enters the regenerator at a point below a grid plate 42 which serves to uniformly distribute the air over the cross-sectional area of the vessel.

By contacting the air with spent catalyst in the regenerator the temperature of the solids increases from the combustion of the contaminating material. In the case of using chromium oxide supported on alumina, such as in this example, the temperature of regeneration is maintained at a temperature of 1100 F. The temperature is maintained at this point by means of a plurality of vertical tubes 44 which are directly on contact with the catalyst bed in the regenerator. The heat of combustion is removed by heated water which is supplied from a line 45 into a boiler drum 46. The water flows from the boiler drum through a line 47 into a common header 48 which is located outside of the regenerator and serves to distribute same through the vertical tubes which are disposed in the catalyst bed. The heated water and stem which is formed in tubes 44 flow from the vertical tubes into a second common header 49 located outside of the regenerator and through a line 50 into drum 46, wherein the steam is separated and removed through an overhead line 51. Ordinarily, the rate of combustion air which is fed to the regenerator is regulated to produce superficial linear gas velocities in the order of about 0.1 to about 3 feet per second. The flue gases leaving the catalyst bed in the regenerator usually contain entrained solids, which are separated from the gaseous stream by means of filters 54 which are located in the upper enlarged section 55 of the regenerator. It is possible to employ cyclones for catalyst recovery at this point in the reactor. The separated gases leave the system through lines 56 which are connected to the filters and thence flow through a valved exit line 57. In order to avoid an excessive catalyst deposit on the filters, these filters are blown back by air which is charged through lines 59 and which are connected to lines 56. The blow back gas is supplied to lines 59 by means of line 60. The regenerated catalyst leaves the bottom of the regenerator through a segmental well 62 and thence flows through a line or standpipe 63. This standpipe is, for example, a pipe of 4 inch diameter and contains a slide valve 65 for automatically controlling the flow of solids. The catalyst in this standpipe can have a density of about 10 to pounds per cubic foot or, in this example, 50 pounds per cubic foot. In order to maintain the catalyst in a fluidized condition, hydrogen-containing gas or recycle gas is fed into the standpipe by means of line 66. In order to assist in the upflow of the regenerated catalyst through well 15 hydrogen-containing gas or recycle gas is introduced at the bottom thereof by means of line 67. The main volume of hydrogen-containing gas or recycle gas is fed to the bottom of reactor 10 below grid plate 13 by means of a line 68 and it enters at a temperature of about 1000 F.

Several experiments were conducted on the .pilotplant scale in order to determine the conditions under which reforming and desulfurization should be performed.

Figures 2 and 3 represent a moving and fixed bed system, respectively, wherein experimental tests were conducted.

In Figure 2, the reactor 75 is a vertical cylindrical vessel having a diameter of about 2 inches and a length of about 7 feet. A continuous flow of catalyst was maintained through the reactor by a continuous supply to the top thereof. This was accomplished by introducing catalyst into a hopper 76 and then passing same by means of a screw pump or other suitable means indicated schematically as 77 in the drawing, to a low pressure vessel 78 which is situated therebelow. Catalyst was discharged from the bottom of low pressure vessel 78, by means of a screw pump or other suitable means 79 which was situated below said vessel 78. In order to convey catalyst into high pressure vessel 80, which is disposed below conveying means 79, the passage to the top of vessel 78 was closed; and the catalyst was conveyed from the bottom of vessel into the top of high pressure vessel 80. Any gaseous materials which backflowed into vessel 78 were discharged therefrom by means of a valved vent line 81. The catalyst was introduced into the reactor 75 by a screw pump or suit-able means 87, which interconnected high pressure vessel 80 therewith. Gaseous materials could be vented from vessel 80 by means of vent line 83. Catalyst was discharged from the bottom of reactor 75 in essentially the same manner as described above, except a reverse procedure was used. A conveying means, such as screw pump 85 was connected to the bottom of the re actor. This conveying means was situated above the first discharge vessel or high pressure vessel 86, which also had a conveying means 87 connected to the bottom thereof. In operation, catalyst was removed from the reactor by means of conveying means 85, while the bottom end thereof was sealed off to prevent the escape of any gases from vessel 86, after the desired quantity of catalyst was being removed from the reactor 75. Conveying means 85 was discontinued and any gaseous materials which were present in vessel 86 were discharged through a valved vent line 89. Conveying means 87 was then operated to remove catalyst from the bottom of high pressure vessel 86 and passing same into the top of a low pressure vessel 90. After conveying the catalyst from vessel 86 to vessel 90, conveying means 87 was discontinued and the catalyst was withdrawn from the bottom of vessel 90 by means of a screw pump or conveying means 91. Any gaseous materials which were present in vessel 90 were vented therefrom by means of valved line 92. The catalyst which was discharged from vessel 90 was passed to a regenerator vessel (not shown) wherein the carbonaceous deposits were removed by burning with an oxygencontaining gas, e. g., air.

Naphtha feed was supplied through a line 95 and then transported by means of pump 96 into a suitable pre-heater 97. The naphtha was heated to a temperature of about 400 to about 650 F., and then passed into vaporizer 98, in which a hydrogen-containing gas or recycle gas was passed via line 99. In the vaporizer, the mixture of hydrogen-containing gas and naphtha was further heated to a temperature of about 600 to about 850 F., to insure that the naphtha was vaporized. The mixture of hydrogen-containing gas and naphtha vapor Were then passed from the vaporizer by means of line 99 into a second pro-heater 100. In pre-heater 100, the temperature of the mixture of the vapors and gases was raised to about 850 to about 1200 F. prior to being charged to the bottom part of reactor 75 at a point about 1% feet above the bottom end thereof. After the reactants were in contact with the catalyst for a suitable period of time, the product was discharged from the top of the reactor through an overhead line 103. The product mixture was then passed to condenser 104 which was cooled by means of water. In the condenser, material which is normally liquid at the pressure which exists in the reactor becomes liquefied. The mixture of liquid and gas was then passed from the condenser 104 to a high pressure separator 105. The liquid in the high pressure separator was discharged therefrom through a bottom line 106, which contained a control valve 107, for maintaining the desired pressure in separator 105. The liquid product which was discharged from line 106 entered a flash drum 110, wherein the pressure was maintained at atmospheric in order that the gaseous material absorbed in the liquid product could be flashed off. The gaseous material which flashed off was removed from the flash drum by means of an overhead line 111, and then it passed into a cooler 112, prior to being discharged through a vent line 113. The remaining liquid in flash drum 110 was discharged through a bottom valved line 114 and then it was passed to a product recovery system which is not shown.

The gaseous product material, which remained uncondensed after passing through condenser 104, was discharged from the high pressure separator 105 by means of an overhead line 117. All or a portion of the gas could be vented through a valved line 118 or all or a portion could be passed through a line 119, into which make-up hydrogen was charged by means of line 120. The total gas stream was then passed through a line 121 and thence into a compressor 122 whereby the pressure was raised to about 50 to about 500 p. s. i. g. The gas leaving the compressor was then recycled to vaporizer 98.

In Figure 3, the feed was supplied through a line 125 and then it was conveyed by means of a pump 126 through line 127, wherein hydrogen-containing gas or recycled gas was admixed therewith prior to entering preheater 128. The hydrogen-containing gas was introduced into the line 127 via a valved line 129. The mixture of hydrogen-containing gas and naphtha was heated to an elevated temperature, at which all of the naphtha vaporized, and the total stream was then charged into the top of reactor 132 by means of a line 133. The reactor was a vertical, cylindrical vessel having a diameter of about 2 inches and a length of about 7 feet. This reactor normally contained catalyst to a height of about 6 feet. The reactants passed downwardly through the bed of catalyst, and after being converted in a suitable manner, the product was discharged from the bottom of the reactor via line 134. The reaction product was then passed through a line 135, which led to a condenser 126, which was cooled by means of water. The condenser was maintained at the pressure which existed in the reactor. The material which was normally liquid at reaction pressure was liquefied, and then passed into a separator 137 by means of a line 138, which led from the condenser. Separator 137 was maintained under substantially the same pressure which existed in the reactor. The liquid product in the high pressure separator 137 was discharged through a bottom valved line 139 and then passed into a flash drum 142. The pressure in the flash drum was maintained at atmospheric in order that any gaseous material which was absorbed in the liquid product could be flashed off. The flashed liquid product was then passed to a product recovery system, which is not shown. The gaseous material in the flash drum was discharged from the top thereof through an overhead line 144, and then cooled in a suitable cooling means 145, before being vented through a valved line 146.

The gaseous material which was present in the high pressure separator 137 was discharged therefrom by means of an overhead valved vent line 151. On the other hand, all or a part of this gaseous material could be passed through a line 152 wherein make-up hydrogen was fed by means of a line 153. The total mixture of gas was then compressed in a compressor 155 to a pressure in the order of about 50 to about 500 p. s. i. g., prior to being recycled to the reactor through the recycle line 129.

When the catalyst had obtained a carbon content of about 1 to about 20% by weight, the reaction cycle was discontinued by first stopping the flow of naphtha feed. Hydrogen was continuously passed through the reactor at a temperature of about 500 to about l000 F., and for a period of about 0.1 to about 2 hours, in order that any oil vapor which was present in the reactor may be removed. This constituted the purging cycle. After the reactor was purged for the required period of time, the flow of hydrogen was discontinued. The regeneration cycle was started by passing air through a valved line 156. The air was pro-heated in pre-heater 128, to a temperature of about 500 to about 1000 F., prior to being fed into the top of the reactor via line 133. The flue gases resulting from the regeneration were passed from the bottom of the reactor through line 134, and thence through a valved vent line 158. During the regeneration step, valve 159 in line was kept in a closed position in order to prevent flue gases from passing into the condenser. The regeneration of catalyst was eifected by passing about 3 to about 30 cubic feet of air (at 60 F. and 760 mm.) per pound of catalyst in the reactor. The temperature of regeneration was maintained at about 1000 to about 1300 F.

volumes of naphtha are reformed per volume of catalyst. It is to be particularly noted that the feed stock employed for the experiments was an unusually high sulfur content naphtha. As shown in Table I, the feed stock con- In the pilot plants described in Figures 2 a11d 3, various 5 tains 2.2% sulfur by weight and this material was desulexperiments were conducted for determining the best furized to about 0.1% by weight sulfur by employing cond tions of reformurg and desulfurrzlng of three types not more than about 7 volumes of naphtha per volume of h1gh sulfur-contalnlng naphthas, whrch are shown in of catalyst in the operation. It is also noted that the Table carbon yields for each run shown in Table II are unusu- 10 ally low for a reforming operation on a high-sulfur TABLE I cracked naphtha. Furthermore, good yields of gasoline are obtained along with the hlgh removal of sulfur. The octane improvement is also good, with the exception of Feed Stmk A B C the run in which more than about 7 Volumes of naphtha per volume of catalyst were used. ASTM Distillation: Other types of feed stocks were reformed and desul- IBP 217 218 furized in accordance with the resent invent d 214 238 p 1011 1n or er 249 287 to determine the eflicrency of my process. The results of these experiments are reported in Table III below. as: 2

2 7 313 314 TABLE III 330 335 2%; 2441 1416 1404 1307 29 46.0 46.6 i g 69.4 68.4 I IV Iv IV 89 93 0 6 Pressure, p. s. 1. g....... 200 100 100 100 2 ""6131 Temperature, F. (ave.) 978 862 909 957 30 Space Vel. VO./hT./Ve 0.78 2.02 2.01 0.5 Vol.11aphtha/Vol. Catalyst 4.7 12.12 12.06 3.0 H2 cont. gas feed, cu. ft./bbl. of feed. 6, 040 2, 630 2, 630 2,630 Percent H2 in gas feed 100 100 100 The catalysts employed for this purpose are descrlbed Mols Hg/mo1naphtha. 3.7 b l gea gt n penod, Hrs 6 6 6 6 18 SI vol. percent 100% O4 gasoline 86.9 96.4 05.5 82.2 I601.b peregrtlt Citrate-gasoline. 3.1 011 PGICGII Catalyst Deslgnatlon I H III IV V VI Sulfur, Wt. percent (G4free-gaso1ine) 0.04 0.27 0.13 0, 02 Hz-cu. ft./bbl 2s0 201 127 -227 Octane No. ASTM: M003, Percent 6 18 6 100% G1 gasoline 77.6 68.0 67.7 72.3 0.1203, Percent 20 11 Grime-gasoline 76.6 57.5 67.2 72.7 S101, Wt. Percent.- 5 A1103, Wt. Percent 94 85 82 94 so 89 *Fixed bed operation.

Using the molydena catalyst and the fixed bed oper- It is to be noted from Table III above that in those ation which is shown in Figure 3, the data reported in cases where more than 7 volumes of naphtha per volume Table II was obtained. of catalyst were employed the sulfur content of the C4 TABLE 11* Run N0 1 2 3 4 5 6 F 1 B B B B B B oggalyst II II II II II III Pressure, p. s. 1. g 200 200 100 200 200 200 Temperature, F. (ave) 1, 025 992 949 980 984 980 Space Vel. Vo.[hr./V u 0.40 0.78 1.57 0.30 0.40 0.81 V01. Naphtha/Vol. Catalyst... 2. 4 4. 7 9. 4 1. 8 2. 4 4. 9 Mols Hz/IHOI naphtha 5. 8 5. 6 3. 3. 5 3. 9 4. 112001117. gasfeed, cu.ft./bb1. otfee 5,070 4,890 2,605 6,050 6,025 5,860 Reaction period, hrs 6 6 6 6 6 6 i t 1003 C 1111 e 85 2 92 s 79 5 78 4 84 s 0 81'0911 d gaso G .0

Vol. gercent C4 lice-gasoline" 52.6 76.2 90.3 69.8 72.0 80.1

Carbon, Wt. percen 2.0 0. 7 0. 4 2. 5 2. 4 1. 7

Sulfur, Wt. percent (04 fr 0.05 0.07 .42 003 03 .02

11; yield, cu. ft./bb1 1,150 790 335 420 420 210 N .-AS M. l z at gasoline 86.1 77.6 72.0 82.5 81.0 76.7

C4 free-gasoline 85.0 76.2 71.6 81.5 80.2 76.0

*Fixed bed operation.

The data reported in Table II above was correlated free-gasoline is greater than about 0.1% by weight. As in order to show the relationship between the sulfur conpreviously mdlcated 0.1% by weight of sulfur in C4 tent of the C4 freegasoline and the volume of naphtha free-gasoline corresponds to about 7 volumes of naphtha reformed per volume of catalyst. This relationship is reformed per volume of catalyst. Therefore, it is to be given in Figure 4 of the attached drawings. It is to be noted that the results of using other feed stocks connoted from Figure 4 that the percent of sulfur in the I forms essentially to what has been predlcated from the C4 free-gasoline begins to increase markedly when emdata given m Table II above, 1. e., not more than about ploying more than about 7 volumes of naphtha per vol- 7 volumes of naphtha per volume of catalyst should be ume of catalyst. The increase in sulfur content for an used. operation involving more than about 7 volumes of The chromra-type catalyst was also employed for renaphtha per volume of catalyst is very rapid, thus forrnmg and desulfurrzmg two other types of feed stocks, indicating the unexpected and superior advantages n 85 which are described 1n Table I above. The results of practicing my invention, so that not more than about '7 these experiments are given in Table IV below.

1 1 TABLE IV 12 From Table V above, it is to be noted that all of the experiments were made by processing less than about 7 2332 2364A 1434 1448 volumes of naphtha per volume of catalyst. This condition produced unusual results with respect to an opera- R N 1 2 a 3 a 4 5 tion involving the net consumption of hydrogen. Howgi i B B C 0 ever, it is to be noted that in the case of the data shown Catalyst V V VI VI in Table V, the sulfur content of the C4 free-gasoline 1s g p- -gg---- gig 22 52g 38% greater than 0.1% sulfur which is the figure coinciding ;;g ,f M1 51 (L51 with the 7 volumes of naphtha reformed per volume of Hz'cont. gas feed, cu. itJbbl. of 16 4.920 2, 555 2,620 2,740 10 catalyst. Furthermore, it is to be noted that the carbon c fi 2 i igg 51 8 g production of each of the runs given in Table V above are g f g;,;gj 1 Q; substantially higher than what is produced in the case Vol. naphtha/Vol. catalys 1.9 3.1 1. 2 of the data given in Tables II and IV, under comparable Yields: operating conditions. Therefore, it can be seen that an $2166656256155165: 31 11; $31? 3311 3313 15 operqtionjnvqlving net iz q of y qg n is su Carbon, Wt. percent 1.1 4.3 0.9 0.1 stantially inferior to one in which hydrogen 1S consumed g i i 3 (C41'ree'ga$1m@)- 18 3 93 1 1. with respect to desulfurization and carbon yields. These Octane A two factors are extremely important since a reforming 100% C1 gasoline 75.6 79.3 73.0 and desulfurization operation should be effective and C4 free-585011119 2 5 3 1542 20 economical. The production of high carbon yields causes operating costs to be rather high for the regeneration of Moving bed operation. catalyst. z gg i ggg ggg A comparison was made between molybdena and chromia catalysts, under comparable operating condi- 011 the basls of a reported In Table IV above. tions. These results are given in Table VI below. the per cent desulfurization of feed stock was correlated against the volume of naphtha reformed per volume of catalyst. The per cent desulfurization was determined by TABLE first taking the difference in sulfur content between the C4 free-gasoline and the feed stock and then determining the percentage that this figure constit uted of the sulfur 1397 M34 content of the feed stock. This relationship 18. given n Figure 5 of the attached drawings. On the basis of this F d C C data, it was observed that there is a sharp decrease in dei, Iv VI sulfurization when processing more than about 7 volumes 33 Pressure, p. s. i. g. 100 100 of naphtha per volume of catalyst, in the case of a g f g g 8 g g? chromium oxide catalyst. This finding substantiates the Hgcom ggsiee'd, C f E B 21630 2,520 observations made with respect to the data obtained when Percent H2 in gas feed. 1 1 using a molybdenum ox1de catalyst. It is to be noted g gggf pemdi 6 6 that 111 the case of using more than about 7 volumes of 40 1 percent 100% gasoline 32,2 92,7 naphtha per volume of catalyst, the decrease in per cent V01. percent 01 free-gasoline 79.1 05.0 desulfurization becomes substantially worse than when g Eg g 3 operating within the desired range. 1 ftjbbl f jj j:

For purposes of comparison, reforming and desulfuriza- Octane No. ASTM: tion of the naphtha, containing 2.2% sulfur, was processed gfg ggfi l g under such condit ons that hydrogen was produced. The results are given in Table V below.

TABLE V* Fixed bed operation.

From the above comparison, it is to be noted that the chromia catalyst 18 unexpectedly better than the molyb- Run N0 1 2 3 4 5 Feed Stock B B B B n dena catalyst, with respect to yields of carbon, 100% C4 g y gi g 6 I? 6 gasoline and C4 free-gasoline and the octane numbers of ressur .S. 4 r,- TemDemtmeoF. 940 987 985 1,000 977 .1. the liquid products. While the desulfurization of C4 Space v 1 v 1 v, (181 free-gasoline is slightly better in the case of using the II 01 .i a }1 t 1i 5 3% ggtalystv 1. 32 1. 9; 1 61 1 i3 8 1. 7? molybdena catalyst, the sulfur content of the gasoline in O S 2 a H. cont. gas feed, cu. ft./bbl 4,890 4,910 10,700 2,450 4,680 the @Xpenment Satlsfactory for Catalyst residence time, Hrs 1.6 1.6 1,6 2,1 mercial utilization without further treatment. Hence, Ylelgsll t 7 C 60 the advantages in yields and octane numbers for the $5 5??? Z 8L9 8L9 8L4 84.0 830 chromia catalyst are sufficient in magnitude to justify v01. percent 01 free-gasoselecting it as the preferred catalyst. C g; 2-; 2 g-g Various experiments were made in order to determine gf gf w g g b' 'fig' the effect of hydrogen partial pressure on product yield. gasoline) 0.15 0. 65 0.60 1.51 1. 00 65 The yields were broughtto an equivalent octane basis 7 190 170 220 by first determining the yield-octane number relationship Octane ASTM for the vario b d h d' 100% o, gasoline 74.1 75.5 76.9 76.8 78.5 runs Elven a (We an K 3 Justlng a mimeasoline 73.7 75.0 76.5 76.3 78.1 given yield t0 an equivalent octane basis by means of this relationship. The data obtained are reported in *Moving bed operation. 7 0 Table VII below.

TABLE VII Run N0 1 2 3 4 5 6 7 8 Feed-.- B B B B B B B B Catalyst III III III III III III III III Temperature, 3 F., (ave). 958 984 983 966 963 987 992 980 Pressure, p. s. i. g 100 100 100 100 100 100 200 200 Space Ve1.,Vc./hr./V. 0.80 1.15 0.99 0.82 0.80 1.20 0.78 0.81 Mols Hz/HIOI naphtha. 2.3 1. 27 2.3 3.1 1.6 2.2 5.6 4.4 Hydrogen partial pressure, p. s. 1. a" 40 43 39 38 26 38 183 121 Vol. naphtha/Vol. catalyst 1.28 1.27 1.58 1.31 1.28 1.92 4.7 4.9 Yield: 100C 40OE.P.Gas01ine 79.1 78.7 78.0 78.7 76.9 80.0 86.1 85.2

"All yields reported on a 77 (OFRM) octane basis.

The relationship between hydrogen partial pressure and gasoline yields is given in Fig. 6. It is to be noted that the present process should be operated at a hydrogen partial pressure of at least about 65 p. s. i. a. in order to obtain unusually high yields of gasoline. This characteristic of the process holds true regardless of the other process conditions, because the data in Table IV illustrates that the other conditions can be varied to an extent which is considered significant without disguising the influence of hydrogen partial pressure on yield of product.

Having thus described my invention by furnishing specific examples thereof, it should be understood that no undue limitations and restrictions are to be imposed by reason thereof, but that the scope of my invention is defined by the appended claims.

I claim:

1. A method for reforming and desulfurizing a sulfur containing naphtha consisting essentially of contacting said naphtha containing sulfur in the range of 0.3 to 2.5% by weight and a hydrogen containing gas with a catalyst selected from the group consisting of the oxides of chromium and molybdenum under suitable reforming and desulfurization conditions including a hydrogen partial pressure of at least 50 p. s. i. a., a temperature in the range of 750 to 1150 F., a volumetric space velocity in the range of 0.1 to 4, the relative proportions of naphtha to catalyst being not greater than 7 volumes of naphtha per volume of catalyst and said reaction conditions being selected on the basis of effecting a net consumption of hydrogen and obtaining a product containing not more than 0.1% by weight of sulfur.

2. A method for reforming and desulfurizing a sulfur containing naphtha consisting essentially of contacting said naphtha containing sulfur in the range of 0.3 to 2.5% by weight and a hydrogen containing gas with a catalyst selected from the group consisting of the oxides of chromium and molybdenum under suitable reforming and desulfurization conditions including a hydrogen partial pressure of at least 50 p. s. i. a., a temperature in the range of 750 to 1150 F., a total pressure in the range of 50 to 300 p. s. i. g., a weight space velocity in the range of 0.1 to 4, the relative proportions of naphtha to catalyst being not greater than 7 volumes of naphtha per volume of catalyst, and the reaction conditions being selected on the basis of effecting a net consumption of hydrogen and obtaining a product containing not more than 0.1% by weight of sulfur.

3. A method for reforming and desulfurizing a sulfur containing naphtha consisting essentially of contacting said naphtha containing sulfur in the range of 0.3 to 2.5% by weight and a hydrogen containing gas with a fluid mass of finely divided catalytic particles selected from the group consisting of the oxides of chromium and molybdenum under suitable reforming and desulfurization conditions including a hydrogen partial pressure of at least 65 p. s. i. a., a volumetric space velocity in the range of 0.1 to 4, a temperature in the range of 750 to 1150 F., the total pressure in the range of 75 to 300 p. s. i. g., the relative proportions of naphtha to catalyst being not greater than 7 volumes of naphtha per volume of catalyst, and the reaction conditions being selected on the basis of effecting a net consumption of hydrogen and obtaining a product containing not more than 0.1% by weight of sulfur.

4. A method for reforming and desulfurizing a hydrocarbon oil consisting essentially of contacting a cracked naphtha fraction containing 2.2% by weight of sulfur and a hydrogen containing gas with an oxide of chromium under suitable reforming and desulfurization conditions including a hydrogen partial pressure of at least p. s. i. a., a temperature in the range of 900 to 1025 F., a volumetric space velocity in the range of 0.1 to 1.5, a total pressure in the range of to 300 p. s. i. g., the relative proportions of naphtha to catalyst being not greater than 7 volumes of naphtha per volume of oxide of chromium, and the reaction conditions being selected on the basis of effecting a net consumption of hydrogen and obtaining a product containing not more than 0.1% by weight of sulfur.

5. A method of reforming and desulfurizing a hydrocarbon oil consisting essentially of contacting a cracked naphtha fraction containing 2.2% by weight of sulfur and a hydrogen containing gas with an oxide of molybdenum under suitable reforming and desulfurization conditions including a hydrogen partial pressure of at least 65 p. s. i. a., a temperature in the range of 850 to 1050 F., a total pressure in the range of 75 to 250 p. s. i. g., a volumetric space velocity in the range of 0.1 to 2.2, the relative proportions of naphtha to catalyst being not greater than 7 volumes of naphtha per volume of oxide of molybdenum, and selecting the reaction conditions on the basis of effecting a net consumption of hydrogen and obtaining a product containing not more than 0.1% by weight of sulfur.

6. A method for reforming and desulfurizing a hydrocarbon oil consisting essentially of contacting a naphtha containing sulfur in the range of 0.91 to 2.2% by weight with a catalyst comprising an oxide of a metal selected from the group consisting of molybdenum and chromium under suitable reforming and desulfurization conditions including a hydrogen partial pressure in the range of 65 to 200 p. s. i. a., a temperature in the range of 650 to 1050 F., a volumetric space velocity in the range of 0.1 to 4.0, the relative proportions of naphtha to catalyst being not greater than 7 volumes of naphtha per volume of catalyst and the reaction conditions being selected on the basis of effecting a net consumption of hydrogen and obtaining a product which has an ASTM octane number in the range of 72.7 to and whose sulfur content is not more than 0.1% by weight of sulfur.

References Cited in the file of this patent UNITED STATES PATENTS Number Name Date 2,414,973 Nelson Ian. 28, 1947 2,436,340 Upham et a1 Feb. 17, 1948 2,498,559 Layng et al Feb. 27, 1950 2,500,146 Fleck Mar. 14, 1950 2,528,586 Ford Nov. 7, 1950 2,547,380 Fleck Apr. 3, 1951 2,574,451 Porter et al. Nov. 6, 1951 2,604,436 Adey et al. July 22, 1952 2,604,438 Bannerot July 22, 1952 

1. A METHOD FOR REFORMING AND DESULFURIZING A SULFUR CONTAINING NAPHTHA CONSISTING ESSENTIALLY OF CONTACTING SAID NAPHTHA CONTAINING SULFUR IN THE RANGE OF 0.3 TO 2.5% BY WEIGHT AND A HYDROGEN CONTAINING GAS WITH A CATALYST SELECTED FROM THE GROUP CONSISTING OF THE OXIDES OF CHROMIUM AND MOLYBDENUM UNDER SUITABLE REFORMING AND DESULFURIZATION CONDITIONS INCLUDING A HYDROGEN PARTIAL PRESSURE OF AT LEAST 50 P.S.I.A, A TEMPERATURE IN THE RANGE OF 750* TO 1150* F., A VOLUMETRIC SPACE VELOCITY IN THE RANGE OF 0.1 TO 4, THE RELATIVE PROPORTIONS OF NAPHTHA TO CATALYST BEING NOT GREATER THAN 7 VOLUMES OF NAPHTHA PER VOLUME OF CATALYST AND SAID REACTION CONDITIONS BEING SELECTED ON THE BASIS OF EFFECTING A NET CONSUMPTION OF HYDROGEN AND OBTAINING A PRODUCT CONTAINING NOT MORE THAN 0.1% BY WEIGHT OF SULFUR. 